[0001] The present invention relates to a method in the production of hydrogen peroxide
by alternate reduction and oxidation of alkylated anthraquinones.
[0002] Hydrogen peroxide is produced, in general, by the so-called anthraquinone process
by alternate reduction and oxidation of alkylated anthraquinones dissolved in suitable
organic solvents. The solution of anthraquinones, the so-called working solution,
is first treated with hydrogen gas in the presence of a catalyst in the so-called
hydrogenation stage, whereupon the working solution is passed to the so-called oxidation
stage in which it is contacted with air or oxygen-containing gas to form hydrogen
peroxide. The reaction schedule of these hydrogenation and oxidation stages, respectively,
may be exemplified as follows
R = alkyl, for instance C
2HS8
[0003] The hydrogen peroxide-containing working solution is then usually passed on to a
so-called extraction stage in which the hydrogen peroxide is removed from the working
solution by extraction with water, whereupon the working solution is recycled to the
hydrogenation stage for another cyclic run through the reaction stages described above.
[0004] The above-mentioned hydrogenation is an important step of this continuous process
and is fraught with considerable difficulites. On hydrogenation, great demands are
placed on a high and uniform productivity, but also on the selectivity of the reaction
to avoid side reactions preventing the hydrogen peroxide formation or making it more
difficult. The selectivity depends upon a number of factors, int.al. the reaction
degree of the anthraquinones, the hydrogenation temperature, the hydrogen gas pressure,
the catalyst and the flow conditions at the catalyst.
[0005] To achieve satisfactory productivity and selectivity in the hydrogenation reactor,
one aims at establishing good contact between the catalyst and the reactants hydrogen
gas and working solution during a well-defined period of time and at a suitably selected
temperature in the reactor. In most cases, this is achieved by intensely mixing hydrogen
gas, working solution and catalyst in the reactor.
[0006] To achieve this intense mixing, use is made, for example in the methods disclosed
by German patent specification 812,426, British patent specification 718,307 and German
patent specification 1,542,089, of a finely-divided catalyst which is held in suspension.
The difficulty in using so-called suspension catalysts is that the working solution,
before it is passed on to the oxidation stage, must not contain any noble metal catalyst
which is the type of catalyst usually employed in these hydrogenation methods, because
it would result in the decomposition of hydrogen peroxide in the oxidation stage.
The working solution must therefore be filtered through a filter device in or outside
the hydrogenation reactor before it is passed on to the oxidation stage. The filter
device, the so-called primary filter, may be complemented with additional filter devices
to ensure that the entire catalyst is separated.
[0007] The use of such filter devices for separating the catalyst from the working solution
creates a number of difficulties. The catalyst which must be held in suspension to
maintain the hydrogenation capacity, covers the surface of the primary filter, and
this results after some time in a reduced catalyst concentration in the hydrogenation
reactor, and this in turn causes a reduction in productivity. Furthermore, the catalyst
particles are crushed by the agitator proper to smaller particles, which means that
the filters will be clogged more quickly as time passes, and this again causes a reduced
flow of working solution through the hydrogenation reactor.
[0008] The catalyst concentration in the hydrogenation reactor can be restored to a tolerable
extent, and also the filtration capacity can by partially restored by periodically
effecting so-called backwashing of the primary filters. By this technique, a part
of the solution which is free from catalyst and which has already left the reactor,
is periodically recycled through the filters, usually by means of pressure-increasing
pump means. However, between these backwashings, the filters are again clogged, which
means that part of the amount of catalyst cannot be utilised efficiently. To compensate
for this, an additional amount of catalyst must be supplied, whereby the cost of the
process is increased considerably.
[0009] To avoid the difficulties encountered in suspension catalysts, various suggestions
have been made. Thus, German patent specification 1,064,343 describes an improved
backwashing method, German patent specification 1,272,292 describes the use of charcoal
filters to improve the separation of finely-divided noble metal catalysts, while British
patent specification 718,307 describes a special hydrogenation reactor without agitating
means to reduce the crushing of catalyst particles.
[0010] All of these known methods of carrying the hydrogenation stage of the anthraquinone
process into effect in the presence of suspension catalysts suffer from the disadvantage
that they require expensive primary filter devices for separating the catalyst from
the working solution, which filter devices, in spite of various backwashing procedures,
are clogged after some time. The primary filters must then be exchanged or subjected
to extensive cleaning operations to restore the filtration capacity to its original
level.
[0011] Another possibility of avoiding the difficulties encountered in connection with suspension
catalysts is to use instead a so-called fixed bed catalyst which consists of a bed
of relatively coarse catalyst particles, usually having a diameter of about 1-10 mm.
In such a catalyst, the working solution and finely-divided hydrogen gas are passed
through the particle bed which is held in position by some type of supporting means,
such as a netting.
[0012] To be able efficiently to utilise the expensive catalyst and working solution, it
is endeavoured to optimise the fixed bed catalyst, such that there is obtained a high
selectivity at high production rates and a small bed volume per weight unit of catalyst
metal.
[0013] Considerable difficulties are encountered in the use of fixed bed catalysts because
of the structure which is formed when the individual particles which may be spherical
or have irregular shape, are packed to form a bed through which hydrogen gas and working
solution are to pass. For an efficient utilisation of the catalyst, adequate contact
between the surface of the catalyst material and the reactants.hydrogen and the anthraquinones
dissolved in the working solution is required. Since it is the hydrogen dissolved
in the working solution which reacts on the active catalyst seats, an efficient transfer
of the gaseous hydrogen from the bubbles to the hydrogen dissolved in the working
solution is of great importance to the utilisation of the catalyst, especially since
the catalyst is extremely active.
[0014] To facilitate the dissolution of hydrogen gas in the working solution, Canadian patent
specification 869,919 proposes introducing into the catalyst bed additional inert
packing layers arranged in accordance with the sandwich principle, which results in
a far greater reactor volume and, contingent thereon, an increased volume of expensive
working solution.
[0015] To increase the productivity in the solid bed catalyst it is, of course, also possible
to increase the flow of hydrogen gas through the catalyst bed, but this may cause
wear on the catalyst particles and thus a loss of catalyst, as is pointed out in U.S.
patent specification 2,837,411 which therefore proposes using an extra tank for saturating
the working solution with hydrogen gas at high pressure, whereupon the working solution
saturated with dissolved hydrogen is passed on to the hydrogenation reactor proper
containing the fixed bed. Also this technique is disadvantageous in that it requires
additional equipment and pressure for the saturation operation as well as an additional
volume of working solution to fill the equipment.
[0016] Another disadvantage encountered in using particle fixed beds is mentioned in Swedish
patent specification 382,200, i.e. the tendency to form separate flow channels for
the working solution and the hydrogen gas through the bed, which means that gas and
liquid are separated so that the catalyst cannot be used efficiently. This phenomenon
is usually termed "channeling".
[0017] Furthermore, a particle fixed bed involves the risk of nonuniform packing density
and, thus, different channel structures in different parts of the bed, as a result
of which the flow of hydrogen gas and working solution will be different in different
parts of the bed, and this again may, because of uncontrolled and frequently too long
contact times, cause the permissible temperature and reaction degree of the anthraquinones
to be exceeded locally, resulting in an increased formation of undesired by-products
and precipitates.
[0018] Furthermore, a general disadvantage of particle fixed beds, as compared with the
more fine-grained suspension catalysts, is that expensive catalyst metal coating the
inner parts of the large fixed bed particles is not readily accessible for reaction
purposes, and this means that the catalyst is not efficiently utilised and the selectivity
frequently becomes inadequate.
[0019] In carrying out the hydrogen peroxide method in accordance with the anthraquinone
process in practical and continuous operations, great demands are placed upon constant
reaction conditions throughout the method cycle in respect of working solution flows,
reaction rates and temperatures in the different method stages, especially in the
hydrogenation stage where, for example, the catalyst activity is reduced more or less
rapidly. If a so-called suspension catalyst is used in the hydrogenation stage, it
is comparatively easy to maintain a constant reaction degree by simply conducting
away a quantity of consumed catalyst and, equally simply, to supply fresh catalyst
suspension to the reactor chamber. Although, when use is made of a hydrogenation stage
with the prior art pellet fixed bed devices, the expensive filter devices required
for suspension hydrogenation can partly be dispensed with, there is instead the disadvantage
that the productivity of the catalyst cannot be maintained at a constant level for
longer periods of time. In view hereof, the prior art pellet fixed beds used in the
hydrogenation stage must be replaced relatively frequently by fresh catalyst pellets,
and this usually means that production must be interrupted for emptying the reactor
container.
[0020] It is known that in the purification of exhaust gases, for example automobile exhaust
gases, use can be made of small volumes of so-called monolithic catalysts to avoid
the large pressure drops and clogging problems in the exhaust emission systems which
are encountered with other types of fixed bed catalysts, for example pellets.
[0021] It is the object of this invention to provide a hydrogenation stage in the production
of hydrogen peroxide by the so-called anthraquinone process, a process stage with
gas-liquid-solid phase reaction, use being made of a catalyst which does not require
any primary filtration, but has all of the advantages of the suspension catalyst,
including controlled temperature and reaction conditions in the reactor, and an efficient
catalyst utilisation with a high degree of selectivity. It has surprisingly been found
that the catalyst utilised in the context of this invention also has a very long life.
[0022] The present invention thus comprises a method in the production of hydrogen peroxide
by alternate reduction and oxidation of alkylated anthraquinones which are dissolved
in one or more organic solvents to form a liquid-working solution which is subjected
to catalytic hydrogenation in that it is contacted with hydrogen in the presence of
a hydrogenation catalyst. In this method, hydrogenation is achieved by causing the
working solution and hydrogen to flow through a catalyst bed consisting of one or
more ordered catalyst bodies built up of a thin-walled and coherent structure with
parallel through channels where the catalyst bodies together form a catalyst bed also
having parallel and equally long channels.
[0023] The catalyst body preferably is built up of alternately planar and corrugated layers
forming bundles of parallel channels, but other embodiments may also be used.
[0024] Furthermore, the catalyst body may be built up of a solid, preferably inert structure
to the walls of which the hydrogenation catalyst is fixed in the form of a thin layer.
The catalyst may be applied to the structure by means of a porous carrier.
[0025] Alternatively, the walls of the catalyst body may be completely built up of a porous
carrier, the pores or lattice of which contains the hydrogenation catalyst. To ensure
stability, the walls preferably should be reinforced with an inert thin woven fibrous
material, such as glass cloth.
[0026] The catalyst body thus comprises a coherent structure which consists of a multiplicity
of parallel through channels, the open diameter of said channels being for example
0.5-10 mm, preferably 1-2 mm. The wall thickness of the structure may lie between
about 0.03 and 1 mm, preferably between 0.1 and 0.3 mm.
[0027] In the method according to the invention, the working solution and the hydrogen dissolved
therein are caused to flow through the parallel channels, to the walls of which the
catalytically active material has been applied, anthrahydroquinone being formed by
the reaction between the hydrogen and the anthraquinones dissolved in the working
solution.
[0028] An important advantage of the present invention is that, because all of the channels
in the catalyst bed are of equal length and substantially geometrically uniform, whereby
a uniformly distributed pressure drop across the channels is ensured, the contact
time can be maintained constant for all partial streams in the catalyst bed. In this
manner, the reaction of anthraquinones in each cross-section of the catalyst bed is
maintained constant, and this in turn has a considerable positive influence upon the
selectivity in the anthraquinone process.
[0029] The catalytically active material may be any of the materials usually employed in
the anthraquinone process, for example noble metals, alone or in mixture. Conventional
catalysts are palladium, platinum, rhodium, or mixtures thereof.
[0030] The walls in the catalyst structure may consist of a non-porous or slightly porous
catalytically inert material which can be directly coated with the catalytically active
noble metal according to known technique. Suitable inert structure materials are glass
cloth, aluminum, or a ceramic material.
[0031] To be able to utilise the structure still more efficiently, it is preferably first
coated with a layer of a porous catalyst carrier to which the hydrogenation catalyst
is then applied. Suitable catalyst carriers are, for example, silicate, silica, alumina
and active carbon which can be applied to the structure in a layer of the desired
thickness, preferably a relatively thin layer.
[0032] It is known that the catalytically active metal preferably should be located in the
outer layer of the catalyst structure where it can be utilised more efficiently. Thus,
British patent specification 1,267,794 discloses a spherical shell catalyst for use
in connection with fixed bed hydrogenation of anthraquinones.
[0033] By coating a coherent and thin structure used in the context of this invention with
a layer of a catalyst carrier, there is obtained a so-called shell catalyst with all
its advantages which imply not only that the expensive noble metal is utilised more
efficiently, but also that an improved selectivity is achieved. Thus, the reaction
degree of the anthraquinones in the catalyst pore system a short distance from the
catalyst surface can be maintained at a level such that the risk of undesired by-products
being formed can be reduced.
[0034] The catalyst bed preferably is built up of one or more cylindrical segments which
may have a height of from 0.1 to about 1 m. A suitable segment height is 0.2 m. The
diameter of the segment may be practically the same as the inner diameter of the hydrogenation
reactor. Several segments can be connected in series in the reactor until there is
obtained a bed volume having sufficient production capacity.
[0035] In the hydrogenation reaction, the working solution and the hydrogen gas are conducted
either countercurrently or cocurrently through the catalyst bed, preferably cocurrently.
A part of the working solution is then recycled and mixed with fresh nonhydrogenated
working solution and fresh hydrogen gas before it is again passed through the catalyst
bed. Recycling of the working solution preferably is effected by means of a pump,
but can also be done by means of a strong flow of hydrogen gas bubbles introduced
at the bottom of the bed, whereby the working solution is transported in the direction
of the rising bubbles (the mammoth pump principle). In the latter case, the catalyst
bed need not completely fill out the diameter of the hydrogenation reactor, whereby
the working solution can be recycled in a downward direction through the free gap
between the inner cylindrical wall of the reactor and the fixed bed cylinder. However,
the method in which the working solution is recycled is of no critical importance
to the concept of the invention.
[0036] The hydrogenation reactor preferably consists of a vertical cylindrical tank having
a removable top wall. The catalyst bed previously prepared is placed within the tank,
either in a single volume or divided into several sections, and is fixed with special
fixing means providing a mechanical protection for the catalyst bed upon insertion
into and removal from the reactor and facilitating connection to suitable hoisting
means.
[0037] Since no primary filtration is required, several hydrogenation reactors may be connected
in series in order optimally to utilise the dependence of the productivity and the
selectivity on such factors as, for example, the hydrogen pressure, temperature and
reaction degree.
[0038] In carrying the method according to the invention into effect, the tank is closed
and then filled with working solution, whereupon the working solution and hydrogen
gas can be circulated through the catalyst bed. In the method according to the invention,
the reaction heat can be optionally cooled off either in the recycling flow or in
the inflow to the reactor. During operation, working solution and hydrogen gas are
continuously supplied to the recycling flow of the reactor which, when passing through
the catalyst bed, forms anthrahydroquinones.
[0039] After the passage through the catalyst bed, a partial flow containing hydrogenated
working solution is continuously withdrawn from the reactor and passed on to further
hydrogenation stages or to the oxidation stage of the anthraquinone process where
anthrahydroquinone is oxidised to anthraquinone while forming hydrogen peroxide. The
hydrogen peroxide formed is then washed out of the working solution with water in
a subsequent extraction stage. In principle, the working solution may then be recycled
to the hydrogenation stage for a further passage through the process cycle described
above.
[0040] The relationship between the working solution flow recycled to the catalyst bed and
the net flow to the hydrogenation reactor may be so selected that the hydrogenation
depth and the temperature are maintained at desired levels along the catalyst fixed
bed with due regard to the formation of by-products.
[0041] By-products are anthraquinone compounds which only with great difficulty or not at
all are capable of forming hydrogen peroxide in the subsequent oxidation stage. The
formation of by-products can be minimised either by using a selected catalyst or by
choosing suitable reaction conditions in respect of temperature and reaction degree
of anthraquinone (the reaction degree is also termed hydrogenation depth). Generally,
low hydrogenation temperatures give a low formation of by-products, but a low hydrogenation
rate and, thus, a low production capacity in the reactor. A high temperature, on the
other hand, gives a high reactive capacity, it is true, but also a greater amount
of by-products upon each passage through the reactor. It therefore is important to
carefully check the temperature and the reaction degree during the reaction.
[0042] The hydrogenation stage is carried out at 40-70°C, generally at about 50-55 C, and
at a reaction rate of up to 80% of the amount of anthraquinone supplied. In this manner
a suitable compromise is obtained between the demand for high production capacity
per unit of volume of the reactor and the amount of by-products formed.
[0043] Depending upon the relationship between the reactor inflow and the flow recycled
to the catalyst bed, the temperature will rise to a greater or less extent when the
working solution and the hydrogen gas pass through the catalyst bed. At the relatively
low recirculation rates used in prior art particle fixed bed reactors, it is necessary,
because of the flow resistance and the risk of wear on the catalytically active material,
to cool the working solution before it is recycled to the fixed bed so that the maximum
permissible temperature is not exceeded during passage through the bed. As a result,
the reaction yield will be lower and the catalyst in part of the catalyst bed will
not be fully utilised.
[0044] One of the advantages offered by the catalyst structure employed in the context of
this invention is that it can operate at high flow loads per unit of surface of the
cross-sectional area of the catalyst bed, without causing large pressure drops that
must be overcome by high pump effects. With these high flow loads there is obtained,
together with dispersed hydrogen gas, a high mixing effect in the channels and thus
adequate contact between the reactants and the catalyst, which has a favourable effect
both on the productivity and the selectivity.
[0045] A further advantage of the fixed bed catalyst according to the present invention
is that the low pressure drops permit such high flow rates within the system that,
if the flow is conducted cocurrently, hydrogen gas is readily recycled to the bed,
and the catalyst system can operate at hydrogen gas loads which, on an average, are
higher than in other fixed bed methods, which in some cases has an effect on productivity
and further improves the flow distribution in the catalyst bed.
[0046] The main advantage achieved by using the catalyst bed structure herein defined resides,
however, in the well-defined and well-controlled reaction conditions which are obtained
by dividing the inflow of working solution and hydrogen gas to the catalyst bed into
a multiplicity of approximately equally large partial flows, the contact times of
which with the catalyst will be of the same order during passage through the bed.
[0047] The method according to the present invention can be carried out under the pressure
conditions usually employed in the hydrogenation stage of the anthraquinone process,
i.e. at an excess pressure of about 10-1000 kPa, preferably 200-390 kPa.
[0048] The accompanying drawing illustrates a flow diagram for an embodiment of the method
according to the present invention.
[0049] In the drawing, a catalyst bed 1 is provided in a plant for continuous hydrogenation
of working solution. The catalyst bed 1 comprises a number of catalyst bodies (segments)
2 which are arranged in series and such that parallel flow channels are formed along
the entire bed 1 in the direction of the arrow 3. The working solution flows through
the catalyst bed 1 in the direction of the arrow 4, and hydrogen gas bubbles in the
direction of the arrow 5. Hydrogen gas is introduced at 6 into the system, and working
solution and hydrogen gas are circulated through the system by means of a circulation
pump 7. Nonhydrogenated working solution is introduced at 8, and hydrogenated working
solution is withdrawn at 9. Some hydrogenated working solution is recycled at 10.
[0050] To further illustrate the present invention, the following Examples are given which
merely are examples of forms of the invention and are not intended to restrict the
scope of the invention as defined by the appended claims.
EXAMPLE 1
Test A - Comparison Test
[0051] In a pilot plant which comprised all of the partial stages for cyclically carrying
the anthraquinone process into effect, the hydrogenation stage was carried out in
a simple vertical tubular reactor loop having a diameter of 100 mm, into which hydrogen
gas was introduced at one leg through a so-called Poral filter in such a manner that
very small bubbles were obtained which, together with the working solution, were circulated
in the loop by means of a pump. By the circulatory movement, the working solution
and the hydrogen gas were conducted downwardly in one leg and upwardly in the other
leg or, if necessary, vice versa in that the direction of rotation of the pump was
reversed.
[0052] The volume in the reactor loop was maintained constant at about 30 litres. The inflow
to the reactor which was the same as the outflow of the working solution, was about
20 litres/hour. During the hydrogenation reaction, an excess pressure of 200 kPa and
a temperature of 55°C were maintained.
[0053] The working solution which was circulated in the pilot plant, consisted of 83 g/l
of THEAK (tetrahydroethyl anthraquinone) -and 55 g/l of EAK (ethyl anthraquinone)
dissolved in a solvent mixture containing 30% by volume of trioctyl phosphate and
70% of "Shellsol AB".
[0054] To the hydrogenation apparatus described above 50 g of catalyst in the form of finely-divided
Pd, so-called palladium black, was batched and held in suspension by circulation pumping
together with hydrogen gas and working solution.
[0055] In the reactor, anthraquinone reacted to anthrahydroquinone, and the hydrogenated
working solution was conducted through a filter candle provided in one leg, before
it was passed on to the oxidation stage. The filter device which consisted of porous
carbon, was backwashed regularly to prevent clogging of the filters.
[0056] During the test which was conducted for about 1000 hours, the reaction of anthraquinones
to anthrahydroquinones was maintained at 45%.
[0057] To maintain the reaction constant at the desired level, further fresh active palladium
catalyst was periodically added during the test. The amount of palladium black added
during 1000 hours was 22 g, while the amount of consumed catalyst withdrawn during
the same period of time was determined at 20 g. In this manner, the amount of catalyst
in the reactor was maintained constant.
[0058] The productivity, calculated in g of H
20
2 produced per g of Pd and hour, was determined at 16 g.
[0059] The selectivity which may be expressed as the loss of active anthraquinone per amount
of H
20
2 produced, was determined at 0.18 mol per 1000 mols of H
20
2 pro- duced.
Test B
[0060] In the pilot plant according to test A, the hydrogenation stage was so modified that
the filter device used in test A was removed and replaced by a tubular segment containing
a fixed bed catalyst according to the invention. The fixed bed catalyst had a diameter
of 80 mm and a volume of 1 litre, and its structure consisted of thin-walled glass
fibre matrix material arranged as alternately planar and corrugated layers forming
parallel bundles of through channels in the direction of flow within the reactor.
The free channel diameter varied between 1.5 and 2 mm. Applied to the structure walls
was a thin layer of porous silica gel, a large surface area of which was coated with
Pd.
[0061] The test.started with the same composition of the working solution and the same flow
of working solution to and from the reactor as at the start of test A.
[0062] The reaction of anthraquinone to anthrahydroquinone was determined at 60% at the
start of the test. The circulation flow in the reactor loop through the catalyst fixed
bed according to the invention was held at a level as high as 300 litres/min to test
the stability of the catalyst. During the major part of the test, the flow of working
solution and hydrogen gas was from the bottom upwards.
[0063] At the beginning of the test, the productivity was determined at 80 g of H
20
2 per g of Pd and hour, or
533 g of H
20
2 per kg of catalyst and hour. A reversal of the flow direction through the bed (from
the top downwards) did not affect the productivity.
[0064] After 1250 hours of operation, the productivity could be determined at 97% of the
value at the beginning of the test.
[0065] The loss of active quinone during the test period was determined at 0.10 mol per
1000 mols of H
20
2 produced.
[0066] The hydrogenated working solution which left the hydrogenation reactor was first
conducted through a very fine-pored filter before entering the oxidation stage. After
1250 hours of operation, it was found that the filter contained about 0.2% of the
catalyst material supplied to the reactor at the beginning of the test, in spite of
the fact that the flow load on the catalyst bed was much higher than was necessary
in order to maintain the productivity.
Test C= Comparison Test
[0067] In the hydrogenation apparatus previously described, the tubular segment containing
the fixed bed catalyst was removed after test B and replaced by a tubular segment
having an inner diameter of 80 mm and provided at its bottom with a netting. The tubular
segment was filled with spherical catalyst pellets having a diameter of about 2 mm
and consisting of a porous alumina carrier with 0.3% by weight of Pd to a level of
100 mm. During the hydrogenation test, working solution and hydrogen gas were circulated
from the top downwards through the bed. The circulation flow in the reactor loop was
reduced as compared with test B to 30 litres/min. because of the higher flow resistance
of the particle fixed bed. The initial composition of the working solution was the
same as in test A. Temperature and pressure were held at the same level as in tests
A and B, and so was the supply of working solution.
[0068] At the beginning of the test, the productivity was measured at 314 g of H
20
2 per kg of catalyst and hour, or 105 g of H
20
2 per g of Pd and hour.
[0069] After about 600 hours of hydrogenation, the productivity had decreased to 90% of
the initial activity, and after 1000 hours to 85%.
[0070] The selectivity expressed as loss of active quinone per 1000 mols of H
20
2 produced was determined at 0.25 mol after 1000 hours.
[0071] As in test B, the hydrogenated working solution was filtered through a fine-pore
filter before entering the oxidation stage. The amount of catalyst mass on the filter,
which had detached itself from the bed, was determined after 1000 hours of continuous
hydrogenation at about 0.5% of the amount of catalyst initially supplied.
EXAMPLE 2
[0072] Two catalyst bodies according to the invention were placed in a pilot reactor having
a constant operating volume of about 270 litres and constructed basically like the
hydrogenation apparatus of Example 1. Each circular segment had a diameter of 342
mm and a height of 200 mm, and the segments were placed in series. The catalyst bodies
covered the entire tubular cross-section of one leg in the tubular reactor loop. Underneath
the catalyst bed, the hydrogen gas was introduced through a Poral filter into the
passing flow of working solution whose circulation rate in the reactor loop during
the test was measured at about 80 m
3/hour. The very large recycling flow which was tenfold greater than was necessary
on productivity grounds, was selected to provide maximum wear load on the catalyst
during the test.
[0073] The inflow of working solution was maintained at about 300 litres/hour during the
test. The working solution used during the test had approximately the same composition
as at the beginning of test A in Example 1. During the test, the hydrogenation temperature
was maintained at about 54°C and the pressure at an excess pressure of about 230 kPa.
[0074] The reaction of anthraquinone to anthrahydroquinone was determined by analysis at
about 50% at the beginning of the test, at a productivity of 67.0 g of H
20
2 per g of Pd and hour, or 390 g per-kg of catalyst bed and hour.
[0075] After more than 5000 hours of operation with these catalyst beds, the productivity
was determined at 98% of the inital value, which shows the surprisingly high stability
of the catalyst as compared with prior art hydrogenation catalysts used in the production
of hydrogen peroxide.
[0076] The hydrogenated working solution which left the reactor device, was passed through
a fine-pore filter for trapping any catalyst mass that might have detached itself
from the bed during the tests.
[0077] The amount of catalyst mass that had detached itself during 5000 hours of operation
was determined at but 0.12% by weight of the originally supplied catalyst, in spite
of the fact that the flow load on the catalyst bed was very high.
EXAMPLE 3
Test A - Comparison Test
[0078] The hydrogenation reactor of the pilot plant for the production of hydrogen peroxide
which was used during the tests in Example 1 was modified in such a way that the tubular
segment with the fixed bed catalyst was dismounted from the tubular reactor loop and
a primary filter device containing a Grade 5 Poral filter was mounted.
[0079] In the pilot plant thus modified, a working solution consisting of 74 g of amyl anthraquinone,
42 g of tetra- amyl anthraquinone, 14 g of ethyl anthraquinone and 24 g of tetraethyl
anthraquinone dissolved in a mixture of the solvents "Shellsol":nonanol:octanol in
a volume ratio of 50:30:20, was circulated.
[0080] The hydrogenation reaction was carried out at 56
QC and at an excess pressure of 200 kPa. The catalyst was a known slurry catalyst which
is used in processes for the production of hydrogen peroxide. The catalyst consisted
of spherical, porous silicate particles having a diameter of 80 um and containing
2% by weight of Pd. At the beginning of the test, about 90 g of catalyst were batched
into the reactor..Productivity was determined at 68 g of H
20
2 per g of Pd and hour at a 31% reaction of anthraquinone to anthrahydroquinone.
[0081] After about 192 hours of operation, productivity decreased to 56 g of H
20
2 per g of Pd and hour.
[0082] Further catalyst was batched so that the amount of catalyst in the reactor was 140
g in all. The productivity then rose temporarily to 74 g H
20
2 per g of Pd and hour at a 53% reaction of anthraquinone to anthrahydroquinone.
[0083] After further operation for 480 hours, productivity had descreased to 29 g of H
20
2 per g of Pd and hour.
[0084] After a test period of 672 hours, the working solution was analysed and the loss
of so-called active quinones was determined at 2.4 mols of quinone per 1000 mols of
H20
2 produced.
[0085] The primary filter was replaced once during the test period. The filter had been
clogged by crushed catalyst particles, in spite of periodic backwashings. Test B
[0086] After test A was finished, the hydrogenation reactor was again modified in that a
tubular segment containing a fixed bed catalyst according to the invention having
a diameter of 80 mm and covering the entire inner diameter of the tube, and having
a height of 140 mm, was substituted for the filter device so that the channels of
the structure were ordered along the direction of the circulatory flow in the reactor.
[0087] In this test, the basic structures consisted of a thin-walled ceramic material having
a geometry resembling the glass fibre structure described in Example 1, test B. The
walls of the ceramic structure were coated with a thin layer having a thickness of
between 60 and 80 um of the catalytically active material which consisted of porous
active alumina containing Pd. During the test, the temperature was maintained at 56°C
and the pressure at an excess pressure of about 200 kPa. The circulatory flow in the
reactor was determined at about 300 litres/hour during the entire test. The flow of
working solution to the reactor was, as in test A, maintained constant at 20 litres/hour
during the test period which lasted for 432 hours. The composition of the working
solution at the beginning of the test resembled the composition at the beginning of
test A. After 432 hours, the activity of the catalyst bed had decreased to 98% of
the initial value which was 64.3 g of H
20
2 per
g of Pd and hour at a reaction of anthraquinone to anthrahydroquinone of about 45%.
[0088] The loss of so-called active quinone was determined at 0.8 mol per 1000 mols of hydrogen
peroxide produced. Test C
[0089] After test B had been finished, the tubular segment containing the fixed bed catalyst
according to test B was replaced by an identical tubular segment containing a fixed
bed catalyst according to the invention in which the basic structure consisted of
thin-walled aluminum sheet metal having a geometry similar to that of test B. The
walls of the aluminum structure were coated with a thin layer of porous silica gel
whose large surface was coated with Pd.
[0090] Test C was conducted with freshly supplied working solution according to test A under
otherwise identical conditions as in test B for a period of 760 hours. After that
period, the activity of the catalyst bed had decreased to 90% of the initial value
which was 61.2 g of H
20
2 per g of Pd and hour at a reaction of anthraquinone to anthrahydroquinone of about
48%.
[0091] The loss of so-called active quinone was determined at 0.5 mol per 1000 mols of hydrogen
peroxide produced.
EXAMPLE 4
Test A - Comparison Test
[0092] In a production plant for the production of hydrogen peroxide, comprising two series-connected
agitator reactors each containing about 350 kg of suspension catalyst (so-called Raney
nickel), continuous hydrogenation of working solution was carried out. The volume
of each reactor was 8 m
3, and the inflow of nonhydrogenated working solution, which was the same as the outflow
of hydrogenated working solution, was maintained constant at 15 m
3/hour for each reactor. The hydrogenation was conducted at about 56
0C and an excess pressure of 170 kPa. During the test, the working solution initially
contained 0.39 mol of THEAK (tetrahydroethyl anthraquinone), 0.305 mol of EAK (ethyl
anthraquinone) and 0.104 mol of THAK (tetrahydroanthraquinone) dissolved in each litre
of working solution where the solvents consisted of a mixture of xylene and 2-octanol
in a volume ratio of 1:1. The reaction of anthraquinones to anthrahydroquinones was
maintained constant at 48% of the total quinone content and 77% of the tetrahydroanthraquinone
proportion during the entire test period which comprised about 7500 hours. The loss
of active anthraquinone during the test period was determined at 0.38 mol/1000 mols
of hydrogen peroxide produced. To maintain the anthraquinone composition of the working
solution constant and to compensate for loss of active anthraquinone, EAK and THAK
were added in minor proportions during the test period. Although means for backwashing
had been provided, it was necessary to replace all primary filters 10 times during
the test period because these filters were clogged by the catalyst. Test B
[0093] In a production plant for the production of hydrogen peroxide, the catalyst according
to the invention was used in the hydrogenation stage. The catalyst had been placed
in the reactor in the form of several series-connected circular segments having a
diameter of 0.92 m according to the drawing. The entire flow of working solution and
finely-divided hydrogen gas conducted through the catalyst bed was uniformly distributed
over the entire cross-sectional area of the bed. The major part of the flow which
had passed the catalyst bed from the bottom upwards, was recycled to the bed. A minor
part of the circulation flow that had passed through the bed left the reactor chamber
and was passed on to the oxidation stage. The outflow of hydrogenated working solution
which was the same as the inflow of nonhydrogenated working solution, was maintained
constant at about 12 m
3/hour. The temperature in the reactor loop was about 54
oC, and the pressure about 250 kPa above atmospheric during the test period. The composition
of the working solution supplied to the hydrogenation reactor was determined at the
beginning of the test at 0.210 mol of EAK, 0.026 mol of THEAK, 0.480 mol of AAK (amyl
anthraquinone) and 0.051 mol of THAAK (tetra- hydroamyl anthraquinone) in every litre
of working solution having a solvent composition of 55% by volume of "Shellsol AB",
30% by volume of nonanol, and 15
% by volume of 2-octanol. During the test period which lasted for 2400 hours, the reaction
on the total quinone content of the working solution was maintained at a total of
about 32% of hydroquinone. The reaction of tetrahydroanthraquinones was then almost
complete, and the reaction of anthraquinones EAK and AAK was about 24%, because the
hydroquinone balance had firmly shifted to tetrahydroanthraquinone..
[0094] It is common knowledge.that tetrahydroanthraquinones are more stable to degradation
in the hydrogenation stage than anthraquinones. In view hereof, use is made in the
majority of A/0 processes of working solution compositions where the majority of the
dissolved quinones are tetrahydroanthraquinones. In these "tetra" systems, normally
only about 70-80% of the tetrahydroalkyl anthraquinones are reacted to tetrahydroalkyl
anthrahydroquinone in order to avoid too rapid a degradation to by-products unable
to form hydrogen peroxide in the subsequent process stages. In these "tetra" systems
alkylanthrahydroquinone does not exist in measurable quantities because the balance
has been shifted. In spite of the, to the hydrogenation stage, unfavourable reaction
conditions during the test period, a loss of but 0.065 mol of active quinones per
1000 mols of hydrogen peroxide produced was established.